EP1226103B1 - Nachrüstbares methanol anlage für verfahren von essigsäure - Google Patents

Nachrüstbares methanol anlage für verfahren von essigsäure Download PDF

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Publication number
EP1226103B1
EP1226103B1 EP00972559A EP00972559A EP1226103B1 EP 1226103 B1 EP1226103 B1 EP 1226103B1 EP 00972559 A EP00972559 A EP 00972559A EP 00972559 A EP00972559 A EP 00972559A EP 1226103 B1 EP1226103 B1 EP 1226103B1
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EP
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Prior art keywords
methanol
hydrogen
stream
plant
carbon monoxide
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Revoked
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EP00972559A
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English (en)
French (fr)
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EP1226103A1 (de
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Daniel Marcel Thiebaut
Kenneth Ebennes Vidalin
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Acetex Cyprus Ltd
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Acetex Cyprus Ltd
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Priority claimed from US09/430,888 external-priority patent/US6274096B1/en
Application filed by Acetex Cyprus Ltd filed Critical Acetex Cyprus Ltd
Priority to EP06006456A priority Critical patent/EP1683780B1/de
Priority to EP09001457A priority patent/EP2146166A3/de
Publication of EP1226103A1 publication Critical patent/EP1226103A1/de
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    • C07—ORGANIC CHEMISTRY
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    • C07C67/00—Preparation of carboxylic acid esters
    • C07C67/36—Preparation of carboxylic acid esters by reaction with carbon monoxide or formates
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    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/32—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air
    • C01B3/34—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents
    • C01B3/38—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts
    • C01B3/384—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts the catalyst being continuously externally heated
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    • C01C1/04—Preparation of ammonia by synthesis in the gas phase
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    • F25J3/04521—Coupling of the air fractionation unit to an air gas-consuming unit, so-called integrated processes
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Definitions

  • the present invention is directed generally to a process for making synthesis gas from which streams of carbon monoxide and methanol can be obtained for the manufacture of acetic acid, and more particularly to the retrofit of a methanol plant to divert all or a portion of the syngas from the existing methanol synthesis loop to a carbon monoxide separator and to react the methanol from the methanol synthesis loop with the carbon monoxide from the separator in approximately stoichiometric proportions to directly or indirectly make acetic acid.
  • the primary raw materials for acetic acid manufacture are, of course, carbon monoxide and methanol.
  • methanol is imported and carbon monoxide, because of difficulties associated with the transport and storage thereof, is generated in situ, usually by reforming natural gas or another hydrocarbon with steam and/or carbon dioxide.
  • a significant expense for new acetic acid production capacity is the capital cost of the equipment necessary for the carbon monoxide generation. It would be extremely desirable if this capital cost could be largely eliminated or significantly reduced.
  • the present invention addresses a way of modifying an existing unprofitable methanol plant to make it more profitable when methanol prices are low and/or gas prices are high.
  • the present invention involves the discovery that the large capital costs associated with CO generation for a new acetic acid plant can be significantly reduced or largely eliminated by retrofitting an existing methanol or methanol/ammonia plant to make acetic acid.
  • All or part of the syngas is diverted from the MEOH synthesis loop and supplied instead to a separator unit to recover CO 2 , CO and hydrogen, which are advantageously used in various novel ways to produce acetic acid.
  • the recovered CO 2 can be supplied to the reformer to enhance CO production, or to the MeOH synthesis loop to make methanol.
  • the recovered CO is usually supplied to the acetic acid reactor with the methanol to make the acetic acid.
  • the recovered hydrogen can be supplied to the MeOH loop for methanol production, used for the manufacture of ammonia or other products, burned as a fuel, or exported, since the hydrogen is normally produced in excess of the requirements for methanol synthesis in the present invention.
  • the carbon dioxide can be fed into a catalytic reformer to which natural gas and steam (water) are fed. Syngas is formed in the reformer wherein both the natural gas and the carbon dioxide are reformed to produce syngas with a large proportion of carbon monoxide relative to reforming without added carbon dioxide.
  • the CO 2 can be supplied to the MeOH loop, with additional CO from the synthesis gas and/or additional imported CO 2 , for catalytic reaction with hydrogen to make methanol.
  • the syngas can be split into a first part and a second part.
  • the first syngas part is converted to methanol in a conventional methanol synthesis loop that is operated at half of the design capacity of the original plant since less syngas is supplied to it.
  • the second syngas part can be processed to separate out carbon dioxide and carbon monoxide, and the separated carbon dioxide can be fed back into the feed to the reformer to enhance carbon monoxide formation, and/or fed to the MeOH synthesis loop to make methanol.
  • the separated carbon monoxide can then be reacted with the methanol to produce acetic acid or an acetic acid precursor by a conventional process.
  • Separated hydrogen which is generally produced in excess beyond that required for methanol synthesis in the present process, can also be reacted with nitrogen, in a conventional manner, to produce ammonia.
  • a portion of acetic acid that is produced can be reacted in a conventional manner with oxygen and ethylene to form vinyl acetate monomer.
  • the nitrogen for the ammonia process (especially for any added ammonia capacity in a retrofit of an original methanol plant comprising an ammonia synthesis loop) and the oxygen for the vinyl acetate monomer process, can be obtained from a conventional air separation unit.
  • the present invention provides, in one aspect, a method for retrofitting an original methanol plant which has at least one steam reformer for converting a hydrocarbon to a syngas stream containing hydrogen and carbon monoxide, a heat recovery section for cooling the syngas stream, a compression unit for compressing the syngas stream, and a methanol synthesis loop for converting at least a portion of the hydrogen and carbon monoxide in the syngas stream to methanol.
  • the method converts the methanol plant into a retrofitted plant for manufacturing a product from carbon monoxide and methanol selected from the group consisting of acetic acid, acetic anhydride, methyl formate, methyl acetate and combinations thereof.
  • the method comprises the steps of: (a) diverting a portion of the syngas stream from at least one reformer to a separation unit; (b) operating the methanol synthesis loop with a feed comprising the remaining syngas stream to produce less methanol than the original methanol plant; (c) operating the separation unit to separate the diverted syngas into at least a carbon monoxide-rich stream and a hydrogen-rich stream, wherein the quantity of hydrogen in the hydrogen-rich stream is greater than any net hydrogen production of the original methanol plant; and (d) reacting the carbon manoxide-rich stream from the separation unit with the methanol from the methanol synthesis loop to form the product, wherein the diversion of the syngas stream is balanced for production of the methanol from the methanol synthesis loop and the carbon monoxide-rich stream from the separation unit for stolchlometric conversion to the product.
  • a method for retrofitting a methanol plant comprising the steps of:
  • At least one steam reformer is modified to increase carbon monoxide production in the syngas stream.
  • the syngas stream preferably comprises carbon dioxide, and the separation unit produces a carbon dioxide-rich stream that is preferably recycled to at least one reformer to increase the carbon monoxide production.
  • the reaction step can include the direct catalytic reaction of methanol and carbon monoxide to form acetic acid as in the Mosanto-BP process, for example, or alternatively can comprise the intermediate formation of methyl formate and isomerization of the methyl formate to acetic acid, the intermediate reaction of CO and two moles of methyl alcohol to form methyl acetate and hydrolysis of the methyl acetate to acetic acid and methanol, or the carbonylation of the methyl acetate to form acetic anhydride.
  • the present invention provides a method for retrofitting an original methanol plant that has at least one steam reformer far converting a hydrocarbon/steam feed to a syngas stream containing hydrogen and carbon monoxide, a heat recovery section for cooling the synges stream, a compression unit for compressing the syngas stream, and a methanol synthesis loop for converting at least a portion of the hydrogen and carbon monoxide in the syngas stream to methanol.
  • the retrofitted plant can manufacture a product from carbon monoxide and methanol selected from the group consisting of acetic acid, acetic anhydride, methyl formate, methyl acetate and combinations thereof.
  • the retrofitting method comprises the steps of (a) modifying at least one steam reformer for operation with a feed comprising a relatively increased carbon dioxide content; (b) diverting a portion of the syngas stream from at least one steam reformer to a separation unit; (c) operating the methanol synthesis loop with a feed comprising the remaining syngas stream to produce less methanol than the original methanol plant, (d) operating the separation unit to separate the diverted syngas into a carbon dioxide-rich stream, a carbon monoxide-rich stream and a hydrogen-rich stream; (e) recycling the carbon dioxide-rich stream from the separation unit to at least one modified steam reformer to increase the carbon monoxide formation relative to the original methanol plant and increase the molar ratio of carbon monoxide to hydrogen; (f) reacting the carbon monoxide-rich stream from the separation unit with the methanol from the methanol synthesis loop to form the product, wherein the diversion of the syngas stream is balanced for the production of the methanol from the methanol synthesis loop
  • the modified steam reformer is preferably modified to operate at a higher temperature to enhance the carbon conversion to carbon monoxide.
  • the separation unit can include a solvent absorber and stripper for carbon dioxide recovery, and a cryogenic distillation unit for carbon monoxide and hydrogen recovery.
  • the compression unit preferably has a three-stage compressor, and the syngas stream diversion preferably occurs between the second and third compression stages.
  • the third compressor stage is preferably modified for operation at a lower throughput than the original methanol plant.
  • the recycle loop compressor can also be modified for operation at a lower throughput.
  • the method can also comprise importing a stream of mixed CO/carbon dioxide stream, for example in a 1:2 to 2:1 molar ratio.
  • the imported stream can be supplied to the methanol synthesis loop or to the separation unit, but is preferably supplied to the reformer wherein the carbon dioxide therein is converted to CO.
  • the method can further comprise the step of reacting the hydrogen in the hydrogen-rich stream with nitrogen to make ammonia.
  • the retrofitted plant can use the hydrogen-rich stream from the separation unit as a primary hydrogen source for the ammonia production. With the additional hydrogen available from the syngas, additional ammonia can be produced in the retrofitted plant relative to the original methanol plant.
  • the method can further comprise installing a vinyl acetate monomer unit for reacting a portion of the acetic acid with ethylene and oxygen to make vinyl acetate monomer.
  • An air separation unit can be installed to make the oxygen for the vinyl acetate monomer unit, and the nitrogen produced from the air separation unit preferably matches the nitrogen required for the additional ammonia production.
  • the present invention provides a process for making hydrogen and a product selected from the group consisting of acetic acid, acetic anhydride, methyl formate, methyl acetate and combinations, thereof, from a hydrocarbon via methanol and carbon monoxide which is effected by retrofit of an existing plant.
  • the process comprises the steps of: (a) catalytically reforming the hydrocarbon with steam in the presence of a minor proportion of carbon dioxide to form a syngas containing hydrogen, carbon monoxide, and carbon dioxide having a molar ratio of R ((H 2- CO 2 )/(CO+CO 2 )) from 2.0 to 2.9; (b) recovering heat from the syngas to form a cooled syngas stream; (c) compressing the cooled syngas stream to a separation pressure; (d) diverting a major portion of the compressed syngas to a separation unit; (e) separating the syngas diverted to the separation unit into a carbon-dioxide-rich stream, a carbon monoxide-rich stream and a hydrogen-rich stream; (f) recycling the carbon dioxide-rich stream to the reforming step; (g) further compressing the remaining minor portion of the syngas to a methanol synthesis pressure higher than the separation pressure; (h) operating a methanol synthesis loop to convert the hydrogen and carbon monoxide in the further compressed syngas
  • the process preferably has a molar ratio of carbon dioxide to hydrocarbon comprising natural gas in feed to the reforming step from 0.1 to 0.5 and a ratio of steam to natural gas from 2 to 6.
  • the methanol synthesis loop can be operated below a total maximum capacity of the methanol synthesis loop.
  • the process can further comprise the step of reacting the hydrogen in the hydrogen-rich stream with nitrogen in an ammonia synthesis reactor to make ammonia.
  • the process can also comprise the step of separating air into a nitrogen stream and an oxygen stream and supplying the nitrogen stream to the ammonia synthesis reactor.
  • the process can further comprise the step of supplying the oxygen stream from the air separation unit to a vinyl acetate synthesis reactor, along with a portion of the acetic acid from the carbon monoxide-methanol reaction step, and ethylene, to produce a vinyl acetate monomer stream.
  • the present invention provides a method for retrofitting an original methanol plant, comprising at least one steam reformer for converting a hydrocarbon to a syngas stream containing hydrogen, carbon monoxide, and carbon dioxide, and a methanol synthesis loop for converting hydrogen and carbon monoxide from the syngas stream to methanol, into a retrofitted plant for manufacturing a product from carbon monoxide and methanol selected from the group consisting of acetic acid, acetic anhydride, methyl formate, methyl acetate and combinations thereof.
  • the method includes (1) separating all or part of the syngas stream in a separation unit into respective streams rich in carbon dioxide, carbon monoxide and hydrogen; (2) operating the methanol synthesis loop with a feed comprising (a) carbon dioxide and (b) a portion of the hydrogen-rich stream; and (3) reacting at least a portion of the carbon monoxide-rich stream from the separation unit with methanol from the methanol synthesis loop to form the product.
  • the feed to the methanol synthesis loop can include imported carbon dioxide and/or a portion of the synthesis gas.
  • all of the syngas stream is supplied to the separation step.
  • the amount of the hydrogen-rich stream is generally in excess of the stoichiometric hydrogen required by the methanol synthesis loop.
  • all of the carbon dioxide-rich stream is supplied to the synthesis loop, and all of the carbon monoxide-rich stream to the reaction step.
  • the retrofitting method comprises (1) supplying a major portion of the syngas stream to a separation unit for separating the syngas stream into respective streams rich in carbon dioxide, carbon monoxide and hydrogen, (2) operating the methanol synthesis loop with a feed comprising the carbon-dioxide-rich stream from the separation unit, a minor portion of the syngas stream, and an additional source of carbon dioxide to produce a methanol stream, and (3) reacting the carbon monoxide-rich stream from the separation unit with the methanol stream from the methanol synthesis loop to form the product.
  • the retrofitting method comprises (1) supplying the syngas stream to a separation unit for separating the syngas stream into respective streams rich in carbon dioxide, carbon monoxide and hydrogen, (2) operating the methanol synthesis loop with a feed comprising the carbon-dioxide-rich stream from the separation unit, a portion of the hydrogen-rich stream from the separation unit, a minor portion of the syngas stream, and carbon dioxide from an additional source, to produce a methanol stream, and (3) reacting the carbon monoxide-rich stream from the separation unit with the methanol stream from the methanol synthesis loop in stoichiometric proportions to form the product.
  • the retrofitting method comprises importing a stream of mixed CO/carbon dioxide stream, for example in a 1:2 to 2:1 molar ratio.
  • the imported stream can be supplied to the methanol synthesis loop or to the separation unit, but is preferably supplied to the reformer for conversion of the carbon dioxide to CO.
  • the present invention provides a process for making hydrogen and a product selected from the group consisting of acetic acid, acetic anhydride, methyl formate, methyl acetate and combinations thereof, from a hydrocarbon via intermediate methanol, carbon monoxide, and carbon dioxide.
  • the process includes (1) reforming the hydrocarbon with steam to form a syngas containing hydrogen, carbon monoxide, and carbon dioxide, (2) recovering heat from the syngas to form a cooled syngas stream, (3) compressing the cooled syngas stream to a separation pressure, (4) processing the syngas in a separation unit to separate a carbon monoxide-rich stream from the hydrogen and carbon dioxide, (5) operating a methanol synthesis loop to react a first portion of the hydrogen from the separation unit with the carbon dioxide from the separation unit, and additional carbon dioxide from another source, to obtain a methanol stream, (6) reacting the carbon monoxide-rich stream from the separation unit with the methanol stream from the methanol synthesis loop in stoichiometric proportions to form a product selected from the group consisting of acetic add, acetic anhydride, methyl formate, methyl acetate and combinations thereof.
  • the product comprises acetic acid
  • the reaction step preferably comprises reacting methanol, methyl formate, or a combination thereof in the presence of a reaction mixture comprising carbon monoxide, water, a solvent and a catalyst system comprising at least one halogenated promoter and at least one compound of rhodium, iridium or a combination thereof.
  • the reaction mixture preferably has a water content up to 20 weight percent. Where the reaction step comprises simple carbonylation, the water content in the reaction mixture is more preferably from 14 to 15 weight percent. Where the reaction step comprises low-water carbonylation, the water content in the reaction mixture is more preferably from 2 to 8 weight percent. Where the reaction step comprises methyl formate isomerization or a combination of isomerization and methanol carbonylation, the reaction mixture more preferably contains a nonzero quantity of water up to 2 weight percent.
  • the reaction step is preferably continuous.
  • an original plant which can be retrofitted according to one embodiment of the present invention has an existing conventional steam reformer unit 10, methanol (MeOH) synthesis unit 12 and preferably ammonia synthesis unit 14 wherein hydrogen for the ammonia synthesis unit 14 is taken as purge stream 18 from the methanol synthesis loop.
  • the retrofit of the present invention is generally applicable to any plant that generates and uses synthesis gas to make methanol.
  • reference to "original plant”' shall mean the plant as built and lnduding any intervening modifications prior to the retrofit of the present invention.
  • the reformer unit 10 is typically a fired furnace containing parallel tube banks filled with conventional reforming catalyst such as alumina-supported nickel oxide, for example.
  • the feed to the reformer(s) is any conventional reformer feed such as a lower hydrocarbon, typically naphtha or natural gas.
  • the reformer can be a single-pass reformer or a two-stage reformer, or any other commercially available reformer, such as, for example, a KRES unit available from Kellogg, Brown & Root, as is known to those skilled in the art
  • the reformer effluent of the original methanol plant can contain any conventional H 2 :CO ratio, but is normally close to 2.0 in plants making solely methanol, and higher, e.g.
  • the hydrogen-containing stream is typically obtained as purge stream 16 from the methanol synthesis unit 12 loop which is necessary to keep the level of hydrogen and inerts from building up in the synthesis gas recirculated through the methanol synthesis unit 12.
  • the original plant of Fig. 1 is retrofitted to produce acetic acid (HAC) using the existing reformer 10 and methanol synthesis unit 12 , and keeping any ammonia synthesis unit 14, as shown in Fig. 2.
  • a portion of the effluent 18 from the reformer 10 is diverted from the methanol synthesis unit 12 via line 20 to a new CO 2 removal unit 22.
  • the CO 2 removal unit 22 separates the stream from line 20 into a CO 2 -rich stream 24 and a CO 2 -lean stream 26 using conventional CO 2 separation equipment and methodology such as, for example, absorption-stripping with a solvent such as water, methanol, generally aqueous alkanolamines such as ethanolamine, diethanolamine, methyldiethanolamine and the like, aqueous alkali carbonates such as sodium and potassium, carbonates, and the like.
  • a solvent such as water, methanol, generally aqueous alkanolamines such as ethanolamine, diethanolamine, methyldiethanolamine and the like, aqueous alkali carbonates such as sodium and potassium, carbonates, and the like.
  • a solvent such as water, methanol, generally aqueous alkanolamines such as ethanolamine, diethanolamine, methyldiethanolamine and the like, aqueous alkali carbonates such as sodium and potassium, carbonates, and the like.
  • the CO 2 recovered from the CO 2 removal unit 22 or from another source can be supplied to the reformer 10.
  • Increasing the CO 2 in the feed to the reformer 10 increases the CO content of the effluent 18.
  • the carbon dioxide content of the reformer feed is increased, the share of the carbon in the carbon monoxide in the product synthesis gas 18 that is supplied from the carbon dioxide increases in relative proportion and the share originating from the hydrocarbon decreases. So, for a given CO production rate, the hydrocarbon feed gas requirement is reduced.
  • the effluent 18 from the modified reformer 10 has a molar ratio of hydrogen minus CO 2 to CO plus CO 2 (referred to in the present specification and claims as the "R ratio" (H 2 - CO 2 )/(CO + CO 2 )), which can be optimized for methanol synthesis, preferably within the range from 2.0 to 2.9.
  • the possibility of optimizing the R ratio arises from the discovery that the hydrogen for the ammonia synthesis no longer needs to be obtained as the methanol purge stream 16, but can instead be recovered from the syngas diverted via line 20 as discussed in more detail below.
  • the CO 2 -lean stream 26 contains primarily CO and hydrogen and can be separated in CO separation unit 28 into a CO-rich stream 30 and a hydrogen-rich stream 32.
  • the separation unit 28 can comprise any equipment and/or methodologies for separating the CO/hydrogen mixture into relatively pure CO and hydrogen streams, such as, for example, semi-permeable membranes, cryogenic fractionation, or the like.
  • Cryogenic fractional distillation is preferred, and can include simple partial condensation without any columns, partial condensation with columns, optionally with a pressure swing absorption (PSA) unit and a hydrogen recycle compressor, or methane wash.
  • PSA pressure swing absorption
  • the CO stream 30 preferably contains less than 1000 ppm hydrogen and less than 2 mole percent nitrogen plus methane.
  • the hydrogen stream 32 which is sent to a nitrogen wash unit preferably contains at least 80 mol% hydrogen, more preferably at least 95 mol% hydrogen.
  • a portion of the hydrogen stream 32 is supplied to the existing ammonia synthesis unit 14 in place of the methanol loop purge stream 16.
  • the quantity of hydrogen produced in the stream 32 is generally much larger than the amount previously supplied via line 16. This is due in large part to the fact that less methanol is made in the retrofitted plant, and thus less hydrogen is consumed for methanol synthesis.
  • the additional hydrogen capacity can be used as a fuel supply, or as a raw hydrogen source for another process, such as, for example, increased ammonia conversion.
  • Additional ammonia can be made by supplying a portion of the additional hydrogen to the existing ammonia synthesis reactor 14 where the ammonia conversion capacity can be increased, and/or by installing additional ammonia synthesis unit 33.
  • the increased ammonia capacity can be complemented by the presence of existing ammonia handling, storage and transport facilities which may be able to accommodate the additional ammonia capacity with little or no modification.
  • the methanol synthesis unit 12 is a conventional methanol conversion unit such as, for example, an ICI reactor.
  • the methanol synthesis unit 12 of the retrofitted plant shown in Fig. 2 is essentially the same as in the original plant prior to the retrofit, except that the quantity of methanol produced is lower, preferably half of that of the original plant.
  • the loop recycle compressor (not'shown) is operated at a lower capacity and the purge stream 16 is considerably reduced in quantity.
  • the purge stream 16 is no longer needed for supplying the hydrogen to the ammonia converter 14, since this is now supplied in the retrofitted plant from the hydrogen stream 32 separated directly from the portion of the reformer 10 effluent 18 diverted from the feed to the methanol synthesis unit 12 via line 20.
  • the purge stream 16 can now be used for fuel and/or as a hydrogen source for hydrodesulfurization of the feed to the reformer 10. Since there is no longer any need to pass the excess hydrogen through the methanol synthesis unit 12 for use in the ammonia unit 14, the feed to the methanol synthesis unit 12, i.e. the effluent 18, can be compositionally optimized for more efficient methanol conversion, as described above.
  • methanol synthesis unit 12 can also be desirable to modify the methanol synthesis unit 12, if desired during the retrofit, to include any other modifications which are not present in the original plant but have become conventional and have been developed for methanol synthesis loops since the construction of the original plant and have not previously been incorporated therein.
  • the amount of syngas in the effluent 18 from the reformer 10 which is diverted to C0 2 /CO/H 2 separation is preferably balanced to provide a stoichiometric ratio of methanol and CO to produce acetic acid therefrom in acetic acid synthesis unit 34.
  • the ratio of CO in line 30 and methanol in line 36 is equal or the methanol is produced at a 10-20% molar excess, e.g. a molar ratio from 1.0 to 1.2.
  • a relatively larger quantity (total kg/hr) of the effluent 18 is diverted into line 20, and the remaining minor portion is fed in line 38 to the methanol synthesis unit 12.
  • the acetic acid synthesis unit 34 employs conventional acetic acid manufacturing equipment and methodology well known and/or commercially available to those skilled in the art, such as, for example, from one or more of the acetic acid manufacturing patents mentioned above.
  • a conventional BP/Monsanto process can be employed, or an improved BP/Monsanto process employing BP-Cativa technology (iridium catalyst), Celanese low water technology (rhodium-lithium acetate catalyst), Millenium low water technology (rhodium-phosphor oxides catalyst), Acetex technology (rhodium-iridium catalyst) and/or dual process methanol carbonylation-methyl formate isomerization.
  • the reaction generally comprises reacting methanol, methyl formate, or a combination thereof in the presence of a reaction mixture comprising carbon monoxide, water, a solvent and a catalyst system comprising at least one halogenated promoter and at least one compound of rhodium, iridium or a combination thereof.
  • the reaction mixture preferably has a water content up to 20 weight percent. Where the reaction comprises simple carbonylation, the water content in the reaction mixture is preferably from 14 to 15 weight percent. Where the reaction comprises low-water carbonylation, the water content in the reaction mixture is preferably from 2 to 8 weight percent.
  • the reaction mixture preferably contains a nonzero quantity of water up to 2 weight percent.
  • the reaction is typically continuous.
  • An acetic acid product is obtained via line 40.
  • a portion of the acetic acid from line 40 can be fed to a conventional vinyl, acetate monomer synthesis unit 42 where it is reacted with ethylene via line 44 and oxygen via line 46 to obtain monomer product stream 48.
  • the oxygen in line 46 can be obtained, for example, using a conventional (preferably cryogenic) air separation unit 50 which also produces a nitrogen stream 52 corresponding to the amount of air from line 54 needed for the oxygen in line 46.
  • the amount of air separated can be matched to produce the nitrogen required via line 52 for the additional ammonia capacity added by ammonia synthesis unit 33 as mentioned above.
  • the CO available for acetic acid manufacture is increased by diverting a much larger portion of the syngas from line 18 via line 20 to the CO 2 removal unit 22, than in the embodiment of Fig. 2.
  • This increases the requirement for CO needed for the methanol synthesis unit 12.
  • the CO requirement for the methanol synthesis unit 12 can also be met by supplying CO 2 , although the hydrogen required for methanol synthesis with CO 2 is higher.
  • the CO 2 stream 24 from the CO 2 removal unit 22 is thus supplied to methanol synthesis unit 12 instead of reformer 10 as in the Fig. 2 embodiment.
  • Additional CO or CO 2 , or a mixture of CO/ CO 2 as needed for the methanol synthesis 12 can be supplied from another source (not shown) via line 25.
  • CO 2 is usually easier to produce and transport than CO, and it is also less toxic.
  • Some of the hydrogen obtained in line 32 from the CO separation unit 28 can be supplied via line 35 to the methanol synthesis unit 12. Since the amount of hydrogen in line 32 is generally greater than the amount of hydrogen required for methanol synthesis unit 12, even when using CO 2 as a feedstock in place of CO, the excess hydrogen via line 37 can be exported for another use, such as, for example, ammonia synthesis or combustion as a fuel gas.
  • the embodiment of Fig. 5 can have advantages over that of Fig. 2 in that the retrofit does not require modification of the reformers 10, and the methanol synthesis unit 12 can be operated at essentially design capacity without significantly changing any stream compositions in the methanol synthesis unit 12. Furthermore, the embodiment of Fig. 5 can be operated without excess methanol production, while at the same time producing hydrogen via line 37 for export.
  • Fig. 6 is similar to Fig. 5 except that there is no syngas supplied to the methanol synthesis unit 12, and it is instead supplied in its entirety to the CO 2 removal unit 22. This has the further advantage of maximizing CO and acetic acid production with slightly higher CO 2 import requirements and slightly less hydrogen production via line 37.
  • the retrofit of an existing methanol plant includes its reconfiguration to produce methanol and CO in stoichiometric ratio for acetic acid manufacture while still producing hydrogen at least sufficient for existing ammonia synthesis.
  • the original plant converts 760 MTPD of natural gas in a single-pass reformer with a steam:carbon molar ratio of 2.8 to form 1760 MTPD synthesis gas containing 74.1 mol% hydrogen, 15.5 mol% CO, 7.1 mol% CO 2 , 3.1 mol% methane, 0.2 mol% nitrogen, dry basis, which is added to the methanol synthesis loop.
  • the feed to the methanol synthesis reactor in the original plant is 5600 MTPD of a feed gas comprising 82.2 mol% hydrogen, 9 mol% methane, 4.8 mol% CO, 3 mol% CO 2 , 0.6 mol% nitrogen, and 0.4 mol% MeOH and other constituents, dry basis.
  • the synthesis gas is converted in the original plant to 1560 MTPD crude methanol which is refined by distillation to give 1335 MTPD of refined product.
  • the net loop purge gas (raw hydrogen gas for ammonia synthesis) in the original plant comprises 84.3 mol% hydrogen, 10.5 mol% methane, 2.2 mol% CO, 2 mol% CO 2 , 0.7 mol% nitrogen, 0.4 mol% MeOH and other constituents, dry basis, and is sufficient for ammonia production in the original plant of 500 MTPD.
  • the original plant is retrofitted in accordance with Figs. 3 and 4 to produce 760 MTPD methanol and 606 MTPD CO which can make 1208 MTPD acetic acid, and 331 MTPD hydrogen which can make 1085 MTPD ammonia.
  • the existing reformer 106 is modified to add a new induced draft fan (not shown) and larger steam superheat, mixed feed and feed coils (not shown).
  • the existing third stage synthesis gas compressor 122 is modified for a lower throughput by replacing the inner bundle including rotor and diaphragms.
  • the existing synthesis gas recirculation compressor 126 is similarly modified for a lower flow rate by replacing the inner bundle including rotor and diaphragms.
  • a new aMDEA CO 2 removal unit 128, a new package refrigeration unit 138, a new cold box unit 142 including driers and CO compressor, a new CO 2 compressor 134 to recycle CO 2 to the reformer 106 and new hydrogen compressors 148 to supply hydrogen to the existing ammonia plant (not shown) and a new 600 MTPD ammonia plant (not shown) are added.
  • the methanol plant Prior to and/or during the retrofit (at least while the existing equipment is operational, before shutdown construction), the methanol plant can be operated at a reduced rate, bypassing approximately 27% of the synthesis gas from the discharge of the synthesis gas compressor 122 directly to the existing ammonia plant.
  • Advantages of the retrofit compared to a completely new CO/MeOH plant are the use of existing units and equipment, such as desulfurization, reforming including waste heat recovery, synthesis gas compressor and circulator, etc. Additional advantage is provided by the use of the existing offsite and infrastructure such as steam generation, water treatment, cooling water system, control room and product loading facilities.
  • Fig. 3 shows a block flow diagram for the retrofitted plant
  • natural gas is supplied in line 100 to desulfurizer/saturator unit 102.
  • the existing sulfur removal system is used to remove any H 2 S and organic sulfur from the process feed gas.
  • This gas is mixed with a recycle stream 104 of hydrogen-rich synthesis gas to provide a hydrogen content of 2.4 mol%, and is heated in the desulfurizer heat exchanger (not shown) and in the convection section 116 of the reformer 106 to a temperature of 330°C.
  • the heated gas enters the desulfurizer unit 102 where organic sulfur compounds are first hydrogenated to hydrogen sulfide over a bed of nickel/molybdenum catalyst (not shown). Below the NiMo catalyst is a bed of zinc oxide adsorbent (not shown) in which the hydrogen sulfide is reacted to form zinc sulfide.
  • the desulfurized feed is passed through the existing saturator where the gas is saturated with water from line 108 to reduce the process steam requirement.
  • the gas exiting the saturator in line 110 is mixed with recycle carbon dioxide in line 112 and medium-pressure steam from line 114 so that the mixed gas has 3 moles of equivalent steam per mole of carbon, where equivalent steam is calculated as moles of steam plus 0.6 times moles of carbon dioxide.
  • Addition of CO 2 to the feed gas is a modification of the existing operation of the reformer 106. This produces more CO and balances the synthesis gas composition in effluent stream 115 for more efficient methanol production as described above.
  • the mixed feed is preheated in the modified mixed feed coils (not shown) in the reformer 106 to 515°C.
  • the hot mixed feed is distributed to the reformer 106 catalyst tubes (not shown), passes down through the nickel reforming catalyst, and reacts to form hydrogen, CO and CO 2 .
  • Pressure and temperature at the outlet of the catalyst tubes is 19 bars (absolute) and 880°C.
  • the heat recovery unit 116 is a convection section of the reformer 106 and includes coils for the high pressure steam boiler, high pressure steam superheating (modified in the retrofitted plant), mixed feed preheating (modified in the retrofitted plant), feed gas preheating (modified in the retrofitted plant), and combustion air preheating.
  • modified coils are provided for steam superheating, mixed feed preheating and feed gas preheating services.
  • the reformer 106 includes a new induced draft fan.
  • the old induced draft fan is used as the forced draft fan in the retrofitted plant.
  • the reformer effluent 115 is used to generate steam, preheat boiler feed water, and provide heat in the reboilers for the topping and refining columns in the existing exchangers (not shown).
  • the synthesis gas in line 118 is compressed from 17.3 bars (absolute) to 41.7 bars (absolute) in the first casing of the existing synthesis compressor (first and second stages 120 ) without modification.
  • the gas in discharge line 121 is then split so that 62% goes to the CO/H 2 production via line 130 and the rest is sent via line 123 to the second casing (third stage compressor 122 supplying syngas to line 125 for MeOH synthesis).
  • the third compressor stage 122 of the synthesis gas compressor handles only 40% of the flow of the original plant.
  • This casing is modified with a new inner bundle including rotor and diaphragms.
  • the gas is then cooled in the original third inter-stage cooler (not shown) and water is separated in the original inter-stage separator (not shown).
  • the make-up gas in line 125 is then mixed with recycle gas from line 124 (see Fig. 4) and compressed in recycle circulator 126 to 80 bars (absolute).
  • the circulator 126 will also handle only 61% of the original flow and therefore requires a new inner bundle including rotor and diaphragms.
  • aMDEA CO 2 removal system 128 Sixty-two percent of the gas from the second stage of the synthesis gas compressor 120 is sent via line 130 to a new aMDEA CO 2 removal system 128.
  • This is a single-stage aMDEA system licensed from BASF in which the circulating solution is 40 wt% aMDEA, designed to reduce the carbon dioxide content in the gas stream 130 from 9.7 vol% to 100 ppmv, on a dry basis, in line 132.
  • the absorber (not shown) is operated at 35° to 40°C and 39.5 bars (absolute).
  • the absorber overhead gas enters a knockout drum (not shown) for separation of any entrained solution.
  • the rich solution from the absorber bottom passes through a hydraulic turbine (not shown) for power recovery.
  • the turbine produces power to help drive one of the lean solution pumps (not shown).
  • the solution then enters a stripper (not shown) designed in three sections: a contact cooler on top, a low-pressure (LP) flash section in the middle, and a stripper section in the bottom.
  • the rich solution from the hydraulic turbine enters the LP flash section which promotes CO 2 flashing by pressure reduction.
  • a semi-lean solution pump (not shown) pumps the solution from the bottom of the LP flash section through the lean/semi-lean solution exchanger (not shown) to the top of the stripper section. The exchanger recovers heat from the lean solution leaving the stripper section.
  • the solution leaving the stripper section is reboiled by low-pressure steam in a CO 2 stripper steam reboiler (not shown).
  • the carbon dioxide and steam from the LP flash section is cooled to 35°C in the contact cooler section. This is accomplished by contact with cooling water. Cooled carbon dioxide having a purity of at least 99 vol% on a dry basis is sent via line 133 to new CO 2 compressor 134, a four-stage, motor-driven, integrally geared turbo compressor which discharges the CO 2 into line 112 at 26 bars (absolute) for recycle to the reformer 106 upstream from the mixed feed coil as mentioned previously.
  • the regenerated lean solution is cooled in the lean/semi-lean solution exchanger and then by cooling water. The cooled lean solution is pumped to the top section of the absorber, and a slipstream can be filtered to remove solids. Make-up water is added to the system via line 136.
  • the synthesis gas in line 132 is cooled to 4.4°C in new package refrigeration unit 138 (see Fig. 4) which uses a screw compressor and ammonia as refrigerant.
  • the chilled synthesis gas from the unit 138 is then passed via line 140 into driers/cold box unit 142 where it is dried and separated cryogenically into noncondensed hydrogen stream 144 and CO stream 146.
  • the driers (not shown) are parallel beds packed with molecular sieve, one of which is on line while the other is being regenerated. In the driers, the moisture content of the gas is reduced below 0.5 ppmv and the carbon dioxide content is reduced below 1 ppmv.
  • Each drier can normally operate 12 hours and regeneration with hot (288°C) reject gas from the CO/H 2 plant and cooling takes 6 hours, allowing 6 hours of stand-by.
  • the separation of CO and H 2 in the cold box employs a partial condensation process using two columns (not shown).
  • the dried gas from the driers is cooled and partially liquefied in feed/effluent exchangers (not shown).
  • the liquid is separated out, while the hydrogen product is superheated and expanded in a hydrogen expansion turbine (not shown).
  • Cold gas from the turbine at 19.5 bars (absolute) is re-heated in the feed/effluent exchangers and leaves the cryogenic unit at 19.0 bars (absolute) and 10°C via line 144.
  • the liquid that was separated out, rich in CO, is flashed into a hydrogen reject column (not shown). Flash gas containing mostly hydrogen is taken from the top of the column, and re-heated in the feed/effluent exchangers to the same temperature and pressure as the gas from the turbine with which it is mixed in line 144 for supply to the hydrogen compressor 148.
  • Reboil for the hydrogen reject column is provided by condensing high pressure CO in a reboiler (not shown).
  • the bottom product from the hydrogen reject column now lean in hydrogen but containing excess methane, is flashed into a CO/CH 4 column (not shown) where CH 4 is separated from the CO and exits the column as a liquid bottoms product.
  • the liquid methane is evaporated and heated to ambient in the feed/effluent exchangers and exits the unit 142 at 3.15 bars (absolute) as fuel gas via line 150.
  • the CO from the top of the CO/CH 4 column is heated in the feed/effluent exchangers and compressed in a CO compressor (not shown) into line 146.
  • the CO compressor is also utilized in the heat pump cycle by cooling CO in one of the feed/effluent exchangers, condensing in the reboilers for the hydrogen reject and CO/CH 4 columns and subcooling in another one of the feed/effluent exchangers.
  • Subcooled liquid CO is used as reflux in the CO/CH 4 column and as refrigerant in the feed/effluent exchangers.
  • the evaporated CO is reheated in one of the feed/effluent exchangers before recompression in the CO compressor.
  • the hydrogen compressor unit 148 comprises three parallel reciprocating, non-lubricated compressors which can each compress 50% of the hydrogen produced to 80 bars (absolute) into line 152. Normally two compressors are on line and the third is a spare.
  • the quantity of hydrogen produced in line 152 is sufficient to produce 1084 MTPD ammonia.
  • the existing ammonia plant (not shown) requires just enough hydrogen to make 500 MTPD ammonia, so the rest of the hydrogen product is used to make 584 MTPD ammonia in a new ammonia plant (not shown).
  • the pressure of the hydrogen required for the new ammonia plant can be more or less than 80 bars, so the hydrogen compressor pressure can be adjusted accordingly.
  • the make-up syngas in line 123 (see Fig. 4) is more balanced for MeOH synthesis and has a lower R ratio of 2.1 relative to the original plant.
  • the lower R value is due to CO 2 reforming and results in lower circulation in recycle line 124 and in very little purge from the MeOH loop via line 160.
  • Gas from the discharge of the recycle stage 126 flows through line 162 to original methanol converter 164 and process flows essentially follow the original process flow scheme in line 166, heat recovery and cooling unit 168, line 170, methanol separator 172, recycle line 124, letdown vessel 174, and fuel gas stream 176, at 61-65% of original operating flows:
  • the operation of the MeOH converter 164 was simulated by Kellogg, Brown & Root to estimate the performance of the converter 164 at the new operating conditions and lower capacity.
  • the MeOH product stream 178 from the letdown vessel 176 is 760 MTPD.
  • An existing methanol plant is reconfigured to produce methanol and CO in a stoichiometric ratio for the manufacture of 1,000,000 metric tons/year of acetic acid in accordance with the embodiment of Fig. 5. Assuming 340 days/year of production, this is 2040 kmol/h of acetic acid, which requires 2040 kmol/h MeOH and 2040 kmol/h CO.
  • the existing methanol plant of Example 2 is reconfigured to produce methanol and CO in a stoichiometric ratio for the manufacture of the maximum quantity of acetic acid in accordance with the embodiment of Fig. 6, i.e. by supplying all of the syngas from the reformers 10 to the new CO 2 removal unit 22.
  • the acetic acid production is 2180 kmol/h or 1.07 x 10 6 metric tons/year.

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Claims (33)

  1. Verfahren zum Nachrüsten einer Methanolanlage, umfassend die Schritte:
    Installieren einer Kohlendioxid-, Kohlenmonoxid- und Wasserstofftrenneinheit (22, 28) in Verbindung mit einer ursprünglichen Methanolanlage, umfassend (i) wenigstens einen Reformierer (10) zum Umwandeln eines Kohlenwasserstoffs in einen Synthesegasstrom, der Kohlendioxid, Kohlenmonoxid und Wasserstoff enthält und (ii) einen Methanolsynthesekreislauf (12) zum Umwandeln von Wasserstoff und Kohlenmonoxid aus dem Synthesegasstrom in Methanol;
    Installieren einer Leitung (20), um wenigstens einen Teil des Synthesegasstromes zur Trenneinheit (22, 28) abzuleiten, um alles oder einen Teil des Synthesegasstroms in jeweilige Ströme zu trennen, die reich an Kohlendioxid, Kohlenmonoxid und Wasserstoff sind; Anpassen der ursprünglichen Methanolanlage zur Lieferung wenigstens eines Teils des kohlendioxidreichen Stroms an den Reformierer (10), an den Methanolsynthesekreislauf (12) oder an eine Kombination davon;
    Anpassen des Methanolsynthesekreislaufs (12) zum Betrieb bei vermindertem Durchsatz relativ zur ursprünglichen Methanolanlage; und
    Installieren eines Reaktors (34) zum Reagieren wenigstens eines Teils des kohlenmonoxidreichen Stroms aus der Trenneinheit (22, 28) mit Methanol aus dem Synthesekreislauf (12), um ein Produkt zu bilden, das gewählt ist aus der Gruppe, die besteht aus Essigsäure, Essigsäureanhydrid, Methylformiat, Methylacetat und Kombinationen davon.
  2. Verfahren nach Anspruch 1, wobei der Methanolsynthesekreislauf (12) mit einer Beschickung betrieben wird, die umfaßt (i) Kohlendioxid und (ii) einen Teil des wasserstoffreichen Abstroms aus der Trenneinheit (22, 28).
  3. Verfahren nach Anspruch 2, wobei die Beschickung des Methanolsynthesekreislaufs (12) eingeführtes Kohlendioxid umfaßt.
  4. Verfahren nach Anspruch 2 oder 3, wobei die Beschickung des Methanolsynthesekreislaufs (12) einen Teil des Synthesegases einschließt.
  5. Verfahren nach Anspruch 2 oder 3, wobei der ganze Synthesegasstrom an den Trennschritt geliefert wird.
  6. Verfahren nach einem der Ansprüche 1 bis 5, wobei die Menge von wasserstoffreichem Strom im Überschuß zum stöchiometrisch durch den Methanolsynthesekreislauf (12) erforderlichen Wasserstoff vorliegt.
  7. Verfahren nach einem der Ansprüche 1 bis 6, wobei der ganze kohlendioxidreiche Strom an den Methanolsynthesekreislauf (12) geliefert wird.
  8. Verfahren nach einem der Ansprüche 1 bis 6, wobei der ganze kohlenmonoxidreiche Strom an den Reaktor (34) geliefert wird.
  9. Verfahren nach einem der Ansprüche 1 bis 4 und 6 bis 8, wobei ein Hauptteil des Synthesegasstromes an die Trenneinheit (22, 28) geliefert wird und der Methanolsynthesekreislauf (12) mit einer Beschickung (24) betrieben wird, die den kohlendioxidreichen Strom aus der Trenneinheit (22), einen geringeren Teil Synthesegasstrom und eine zusätzliche Quelle (25) von Kohlendioxid umfaßt, um einen Methanolstrom (36) zu erzeugen.
  10. Verfahren nach einem der Ansprüche 1 bis 4 und 6 bis 9, wobei der Methanolsynthesekreislauf (12) mit einer Beschickung betrieben wird, die kohlendioxidreichen Abstrom aus der Trenneinheit (22), einen Teil eines wasserstoffreichen Stroms (35) aus der Trenneinheit (28), einen Teil des Synthesegasstroms (38) und Kohlendioxid aus einer zusätzlichen Quelle (25) umfaßt.
  11. Verfahren zum Nachrüsten einer Methanolanlage, umfassend die Schritte:
    Installieren einer Kohlenmonoxid- und Wasserstofftrenneinheit (28) in Verbindung mit einer ursprünglichen Methanolanlage, umfassend (i) wenigstens einen Dampfreformierer (106) zum Umwandeln eines Kohlenwasserstoffs in einen Synthesegasstrom, der Wasserstoff und Kohlenmonoxid enthält, (ii) einen Wärmerückgewinnungsabschnitt (116) zum Kühlen des Synthesegasstroms, (iii) eine Kompressionseinheit (120) zum Komprimieren des Synthesegasstroms und (iv) einen Methanolsynthesekreislauf (12) zum Umwandeln wenigstens eines Teils des Wasserstoffs und Kohlenmonoxids in den Synthesegasstrom zu Methanol bei einer ursprünglichen Methanolproduktionsgeschwindigkeit;
    Installieren einer Leitung zum Ableiten eines Teils des Synthesegasstromes stromabwärts von dem wenigstens einen Reformierer (106) zur Trenneinheit (28), um das abgeleitete Synthesegas in wenigstens einen kohlenmonoxidreichen Abstrom und einen wasserstoffreichen Abstrom zu trennen, wobei die Wasserstoffmenge in dem wasserstoffreichen Strom größer als die Nettowasserstofferzeugung der ursprünglichen Methanolanlage ist;
    Anpassen des Methanolsynthesekreislaufs (12) zum Betreiben mit einer Beschickung, welche den verbleibenden Teil des Synthesegasstromes umfaßt, um Methanol bei einer verminderten Geschwindigkeit relativ zur ursprünglichen Methanolproduktionsgeschwindigkeit zu erzeugen;
    Installieren eines Reaktors (34) zum Reagieren des kohlenmonoxidreichen Stroms aus der Trenneinheit (28) mit dem Methanol aus dem Methanolsynthesekreislauf (12), um ein Produkt zu bilden, das aus der Gruppe gewählt ist, die besteht aus Essigsäure, Essigsäureanhydrid, Methylformiat, Methylacetat und Kombinationen davon, wobei die Aufteilung des Synthesegasstromes ausbalanciert wird zur Produktion von Methanol aus dem Methanolsynthesekreislauf (12) und kohlenmonoxidreichen Strom aus der Trenneinheit (28) zur stöchiometrischen Umwandlung des Produkts.
  12. Verfahren nach Anspruch 11, weiterhin umfassend Modifizierung wenigstens eines Dampfreformierers (10), um Kohlenmonoxiderzeugung im Synthesegasstrom zu erhöhen.
  13. Verfahren nach einem der Ansprüche 1 bis 6, 8 und 11 bis 12, wobei der Synthesegasstrom Kohlendioxid umfaßt, der wenigstens eine Reformierer (106) ein katalytischer Dampfreformierer ist, die Trenneinheit (128) einen kohlendioxidreichen Strom (133) erzeugt und wenigstens ein Teil des kohlendioxidreichen Stroms rezykliert wird zu wenigstens einem Dampfreformierer (112), um die Kohlenmonoxiderzeugung zu verstärken.
  14. Verfahren nach Anspruch 12, wobei der Synthesegasstrom in der ursprünglichen Anlage ein Molverhältnis R ((H2-CO2)/(CO+CO2)) kleiner als 2,0 oder größer als 2,9 hat und wobei die nachgerüstete Anlage ein R-Verhältnis von 2,0 bis 2,9 hat.
  15. Verfahren nach einem der Ansprüche 13 oder 14, wobei das Rezyklieren des kohlendioxidreichen Stroms (133) aus der Trenneinheit (128) zu dem wenigstens einen modifizierten Dampfreformierer (106) die Kohlenmonoxidbildung relativ zur Originalmethanolanlage verstärkt und das Molverhältnis von Kohlenmonoxid zu Wasserstoff verstärkt.
  16. Verfahren nach Anspruch 12, wobei der modifizierte Dampfreformierer (106) modifiziert ist, um bei einer höheren Temperatur zu arbeiten.
  17. Verfahren nach einem der Ansprüche 1 bis 16, wobei die Trenneinheit (22 28) einen Lösungsmittelabsorbierer und Abstreifer für Kohlendioxidrückgewinnung und eine kryogene Destillationseinheit für Kohlenmonoxid- und Wasserstoffrückgewinnung umfaßt.
  18. Verfahren nach Anspruch 11 oder 12, wobei die Kompressionseinheit (120, 122) einen Dreistufenkompressor umfaßt und die Synthesegasstromableitung zwischen der zweiten Anlage und dritten Kompressionsstufen auftritt.
  19. Verfahren nach Anspruch 18, weiterhin umfassend Modifizieren der dritten Kompressorstufe (122) zum Betrieb bei einem niedrigeren Durchsatz als die ursprüngliche Methanolanlage.
  20. Verfahren nach Anspruch 11, 12, 16, 18 oder 19, wobei der Methanolsynthesekreislauf (12) der Originalmethanolanlage einen Rezyklierungskreislaufkompressor (134) umfaßt, wobei der Rezyklierungskreislaufkompressor für einen Betrieb bei einem niedrigeren Durchsatz modifiziert ist.
  21. Verfahren nach einem der Ansprüche 1 bis 20, weiterhin umfassend den Schritt des Reagierens des Wasserstoffs in dem wasserstoffreichen Strom (32) mit Stickstoff, um Ammoniak herzustellen.
  22. Verfahren nach einem der Ansprüche 1 bis 19, wobei in der ursprünglichen Methanolanlage ein wasserstoffreicher Abstrom, der einen Kreislaufablaß (16) aus dem Methanolsynthesekreislauf (12) umfaßt, reagiert wurde mit Stickstoff, um Ammoniak zu machen und in der nachgerüsteten Anlage des wasserstoffreichen Stroms aus der Trenneinheit (28) verwendet wird als Hauptwasserstoffquelle zur Ammoniakproduktion.
  23. Verfahren nach Anspruch 22, wobei zusätzlicher Ammoniak in der nachgerüsteten Anlage im Verhältnis zur ursprünglichen Methanolanlage hergestellt wird.
  24. Verfahren nach Anspruch 23, wobei das Produkt Essigsäure umfaßt und das Verfahren weiterhin den Schritt des Einrichtens einer Vinylacetatmonomereinheit (42) zum Reagieren eines Teils der Essigsäure mit Ethylen und Sauerstoff umfaßt, um Vinylacetatmonomer herzustellen.
  25. Verfahren nach Anspruch 24, weiterhin umfassend Einrichten einer Lufttrenneinheit (50), um den Sauerstoff für die Vinylacetatmonomereinheit (42) zu machen und wobei Stickstoff, der aus der Lufttrenneinheit (50) erzeugt ist, sich mit dem Stickstoff mischt, der erforderlich ist für die zusätzliche Ammoniakproduktion.
  26. Verfahren nach einem der Ansprüche 1 bis 25, wobei das Produkt Essigsäure umfaßt und die Reaktionsstufe das Reagieren von Methanol, Methylformiat oder eine Kombination davon in Gegenwart eines Reaktionsgemisches umfaßt, umfassend Kohlenmonoxid, Wasser, ein Lösungsmittel und ein Katalysatorsystem, das wenigstens einen halogenierten Promotor und wenigstens eine Verbindung von Rhodium, Iridium oder einer Kombination davon umfaßt.
  27. Verfahren nach Anspruch 26, wobei das Reaktionsgemisch einen Wassergehalt von bis zu 20 Gewichtsprozent hat.
  28. Verfahren nach Anspruch 27, wobei der Reaktionsschritt einfache Carbonylierung umfaßt und der Wassergehalt in dem Reaktionsgemisch bei 14 bis 15 Gewichtsprozent liegt.
  29. Verfahren nach Anspruch 27, wobei der Reaktionsschritt eine Carbonylierung mit geringem Wassergehalt umfaßt und der Wassergehalt in dem Reaktionsgemisch bei 2 bis 8 Gewichtsprozent liegt.
  30. Verfahren nach Anspruch 27, wobei der Reaktionsschritt Methylformiatisomerisierung oder eine Kombination von dieser Isomerisierung und Methanolcarbonylierung umfaßt und das Reaktionsgemisch eine Wassermenge von bis zu 2 Gewichtsprozent enthält, die nicht Null ist.
  31. Verfahren nach Anspruch 27, wobei der Reaktionsschritt kontinuierlich ist.
  32. Verfahren nach einem der Ansprüche 1 bis 25, wobei der Reaktionsschritt die zwischenzeitliche Bildung von Methylformiat und Isomerisierung des Methylformiats zu Essigsäure umfaßt.
  33. Verfahren nach einem der Ansprüche 1 bis 25, wobei der Reaktionsschritt die Zwischenreaktion von CO und zwei Molen Methylalkohol umfaßt, um Methylacetat zu bilden sowie Hydrolyse von Methylacetat zu Essigsäure und Methanol.
EP00972559A 1999-11-01 2000-10-31 Nachrüstbares methanol anlage für verfahren von essigsäure Revoked EP1226103B1 (de)

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US09/430,888 US6274096B1 (en) 1999-11-01 1999-11-01 Methanol plant retrofit
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US09/547,831 US6232352B1 (en) 1999-11-01 2000-04-12 Methanol plant retrofit for acetic acid manufacture
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PCT/CY2000/000004 WO2001032594A1 (en) 1999-11-01 2000-10-31 Methanol plant retrofit for manufacture of acetic acid

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EP2146166A2 (de) 2010-01-20
WO2001032594A1 (en) 2001-05-10
EP1683780B1 (de) 2009-03-04
RU2002114828A (ru) 2004-01-20
NO328072B1 (no) 2009-11-23
ATE424378T1 (de) 2009-03-15
EP1226103A1 (de) 2002-07-31
ES2263498T3 (es) 2006-12-16
CA2388961A1 (en) 2001-05-10
WO2001032594B1 (en) 2001-10-18
DE60027453T2 (de) 2007-05-03
NO20022063D0 (no) 2002-04-30
AU781369B2 (en) 2005-05-19
AU1127801A (en) 2001-05-14
EP2146166A3 (de) 2012-12-19
MY130262A (en) 2007-06-29
RU2250894C2 (ru) 2005-04-27
ATE323668T1 (de) 2006-05-15
CA2388961C (en) 2011-01-04
NZ519314A (en) 2003-10-31
ES2322910T3 (es) 2009-07-01
EP1683780A1 (de) 2006-07-26
NO20022063L (no) 2002-06-26
PT1683780E (pt) 2009-06-22
US6232352B1 (en) 2001-05-15
DE60041723D1 (de) 2009-04-16

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